Following the introduction of catalytic cracking processes in petroleum refining in the early 1930s, large amounts of olefins, particularly light olefins such as ethylene, propylene, butylene, became available in copious quantities from catalytic cracking plants in refineries. While these olefins may be used as petrochemical feedstock, many conventional petroleum refineries producing petroleum fuels and lubricants are not capable of diverting these materials to petrochemical uses. Processes for producing fuels from these cracking off gases are therefore desirable and from the early days, a number of different processes evolved. The early thermal polymerization process was rapidly displaced by the superior catalytic processes of which there was a number. The first catalytic polymerization process used a sulfuric acid catalyst to polymerize isobutene selectively to dimers which could then be hydrogenated to produce a branched chain octane for blending into aviation fuels. Other processes polymerized isobutylene with normal butylene to form a co-dimer which again results in a high octane, branched chain product. An alternative process uses phosphoric acid as the catalyst, on a solid support and this process can be operated to convert all the C3 and C4 olefins into high octane rating, branched chain polymers. This process may also operate with a C4 olefin feed so as to selectively convert only isobutene or both n-butene and isobutene. This process has the advantage over the sulfuric acid process in that propylene may be polymerized as well as the butenes and at the present time, the solid phosphoric acid [SPA] polymerization process remains the most important refinery polymerization process for the production of motor gasoline.
In the SPA polymerization process, feeds are pretreated to remove hydrogen sulfide and mercaptans which would otherwise enter the product and be unacceptable, both from the view point of the effect on octane and upon the ability of the product to conform to environmental regulations. Typically, a feed is washed with caustic to remove hydrogen sulfide and mercaptans, after which it is washed with water to remove organic basis and any caustic carryover. Because oxygen promotes the deposition of tarry materials on the catalyst, both the feed and wash water are maintained at a low oxygen level. Additional pre-treatments may also be used, depending upon the presence of various contaminants in the feeds. With the most common solid phosphoric acid catalyst, namely phosphoric acid on kieselguhr, the water content of the feed needs to be controlled carefully because if the water content is too high, the catalyst softens and the reactor may plug. Conversely, if the feed is too dry, coke tends to deposit on the catalyst, reducing its activity and increasing the pressure drop across the reactor. As noted by Henckstebeck, the distribution of water between the catalyst and the reactants is a function of temperature and pressure which vary from unit to unit, and for this reason different water concentrations are required in the feeds to different units. Petroleum Processing Principles And Applications, R. J. Hencksterbeck McGraw-Hill, 1959.
There are two general types of units used for the SPA process, based on the reactor type, the unit may be classified as having chamber reactors or tubular reactors. The chamber reactor contains a series of catalyst beds with bed volume increasing from the inlet to the outlet of the reactor, with the most common commercial design having five beds. The catalyst load distribution is designed to control the heat of conversion.
Chamber reactors usually operate with high recycle rates. The recycle stream, depleted in olefin content following polymerization, is used to dilute the olefin at the inlet of the reactor and to quench the inlets of the following beds. Chamber reactors usually operate at pressure of approximately 3500-5500 kPag (about 500-800 psig) and temperature between 180° to 200° C. (about 350°-400° F.). The conversion, per pass of the unit, is determined by the olefin specification in the LPG product stream. Fresh feed LHSV is usually low, approximately 0.4 to 0.8 hr−1. The cycle length for chamber reactors is typically between 2 to 4 months.
The tubular reactor is basically a shell-and-tube heat exchanger in which the polymerization reactions take place in a number of parallel tubes immersed in a cooling medium and filled with the SPA catalyst. Reactor temperature is controlled with the cooling medium, invariably water in commercial units, that is fed on the shell side of the reactor. The heat released from the reactions taking place inside the tubes evaporates the water on the shell side. Temperature profile in a tubular reactor is close to isothermal. Reactor temperature is primarily controlled by means of the shell side water pressure (controls temperature of evaporation) and secondly by the reactor feed temperature. Tubular reactors usually operate at pressure between 5500 and 7500 kPag (800-1100 psig) and temperature of around 200° C. (about 400° F.). Conversion per pass is usually high, around 90 to 93% and the overall conversion is around 95 to 97%. The space velocity in tubular reactors is typically high, e.g., 2 to 3.5 hr−1 LHSV. Cycle length in tubular reactors is normally between 2 to 8 weeks.
For the production of motor gasoline only butene and lighter olefins are employed as feeds to polymerization processes as heavier olefins up to about C10 or C11 can be directly incorporated into the gasoline. With the PSA process, propylene and butylene are satisfactory feedstocks and ethylene may also be included, to produce a copolymer product in the gasoline boiling range. Limited amounts of butadiene may be permissible although this diolefin is undesirable because of its tendency to produce higher molecular weight polymers and to accelerate deposition of coke on the catalyst. The process generally operates under relatively mild conditions, typically between 150° and 200° C., usually at the lower end of this range between 150° and 180° C., when all butenes are polymerized. Higher temperatures may be used when propylene is included in the feed. In a well established commercial SPA polymerization process, the olefin feed together with paraffinic diluent, is fed to the reactor after being preheated by exchange with the reaction effluent. Control of the heat release in the reactor is accomplished in unit with chamber type reactors by feed dilution and recycle quench between the catalyst beds in the reactor and with tubular reactor units, temperature control is achieved by means of the coolant medium surrounding the reactors. The solid phosphoric acid catalyst used is non-corrosive, which permits extensive use of carbon steel throughout the unit. The highest octane product is obtained by using a butene feed, with a product octane rating of [R+M]/2 of 91 being typical. With a mixed propylene/butene feed, product octane is typically about 91 and with propylene as the primary feed component, product octane drops to typically 87.
In spite of the advantages of the SPA polymerization process, which have resulted in over 200 units being built since 1935 for the production of gasoline fuel, a number of disadvantages are encountered, mainly from the nature of the catalyst. Although the catalyst is non-corrosive, so that much of the equipment may be made of carbon steel, it does lead it to a number of drawbacks in operation. First, the catalyst life is relatively short as a result of pellet disintegration which causes an increase in the reactor pressure drop. Second, the spent catalyst encounters difficulties in handling from the environmental point of view, being acidic in nature. Third, operational and quality constraints limit flexible feedstock utilization. Obviously, a catalyst which did not have these disadvantages would offer considerable operating and economic advantages.
In recent years, environmental laws and regulations the have limited the amount of benzene which is permissible in petroleum motor fuels. These regulations have produced substantial changes in refinery operation. To comply with these regulations, some refineries have excluded C6 compounds from reformer feed so as to avoid the production of benzene directly. An alternative approach is to remove the benzene from the reformate after it is formed by means of an aromatics extraction process such as the Sullfolane Process or UDEX Process. Well-integrated refineries with aromatics extraction units associated with petrochemical plants usually have the ability to accommodate the benzene limitations by diverting extracted benzene to petrochemicals uses but it is more difficult to meet the benzene specification for refineries without the petrochemical capability. While sale of the extracted benzene as product to petrochemicals purchasers is often an option, it has the disadvantage of losing product to producers who will add more value to it and, in some cases, transportation may present its own difficulties in dealing with bulk shipping of a chemical classed as a hazardous material.
The removal of benzene is, however, accompanied by a decrease in product octane quality since benzene and other single ring aromatics make a positive contribution to product octane. Certain processes have been proposed for converting the benzene in aromatics-containing refinery streams to the less toxic alkylaromatics such as toluene and ethyl benzene which themselves are desirable as high octane blend components. One process of this type was the Mobil Benzene Reduction (MBR) Process which, like the closely related MOG Process, used a fluidized zeolite catalyst in a riser reactor to alkylate benzene in reformate to from alkylaromatics such as toluene. The MBR and MOG processes are described in U.S. Pat. Nos. 4,827,069; 4,950,387; 4,992,607 and 4,746,762.
Another problem facing petroleum refineries without convenient outlets for petrochemical feedstocks is that of excess light olefins. Following the introduction of catalytic cracking processes in petroleum refining in the early 1930s, large amounts of olefins, particularly light olefins such as ethylene, propylene, butylene, became available in copious quantities from catalytic cracking plants in refineries. While these olefins are highly useful as petrochemical feedstocks, the refineries without petrochemical capability or economically attractive and convenient markets for these olefins may have to use the excess light olefins in fuel gas, at a significant economic loss or, alternatively, convert the olefins to marketable liquid products. A number of different polymerization processes for producing liquid motor fuels from cracking off-gases evolved following the advent of the catalytic cracking process but at the present, the solid phosphoric acid [SPA] polymerization process remains the most important refinery polymerization process for the production of motor gasoline. This process has however, its own drawbacks, firstly in the need to control the water content of the feed closely because although a limited water content is required for catalyst activity, the catalyst softens in the presence of excess water so that the reactor may plug with a solid, stone-like material which is difficult to remove without drilling or other arduous operations. Conversely, if the feed is too dry, coke tends to deposit on the catalyst, reducing its activity and increasing the pressure drop across the reactor. Environmental regulation has also affected the disposal of cracking olefins from these non-integrated refineries by restricting the permissible vapor pressure (usually measured as Reid Vapor Pressure, RVP) of motor gasolines especially in the summer driving season when fuel volatility problems are most noted, potentially creating a need for additional olefin utilization capacity.
Refineries without their own petrochemicals plants or ready markets for benzene or excess light olefins therefore encounter problems from two different directions and for these plants, processes which would enable the excess olefins and the benzene to be converted to marketable products would be desirable.
The fluid bed MBR Process uses a shape selective, metallosilicate catalyst, preferably ZSM-5, to convert benzene to alkylaromatics using olefins from sources such as FCC or coker fuel gas, excess LPG or light FCC naphtha. Normally, the MBR Process has relied upon light olefin as alkylating agent for benzene to produce alkylaromatics, principally in the C7-C8 range. Benzene is converted, and light olefin is also upgraded to gasoline concurrent with an increase in octane value. Conversion of light FCC naphtha olefins also leads to substantial reduction of gasoline olefin content and vapor pressure. The yield-octane uplift of MBR makes it one of the few gasoline reformulation processes that is actually economically beneficial in petroleum refining.
Like the MOG Process, however, the MBR Process required considerable capital expenditure, a factor which did not favor its widespread application in times of tight refining margins. The MBR process also used higher temperatures and C5+ yields and octane ratings could in certain cases be deleteriously affected another factor which did not favor widespread utilization. Other refinery processes have also been proposed to deal with the problems of excess refinery olefins and gasoline; processes of this kind have often functioned by the alkylation of benzene with olefins or other alkylating agents such as methanol to form less toxic alkylaromatic precursors. Exemplary processes of this kind are described in U.S. Pat. Nos. 4,950,823; 4,975,179; 5,414,172; 5,545,788; 5,336,820; 5,491,270 and 5,865,986.
While these known processes are technically attractive they, like the MOG and MBR processes, have encountered the disadvantage of needing to a greater or lesser degree, some capital expenditure, a factor which militates strongly against them in present circumstances.
For these reasons, a refinery process capable of being installed at relatively low capital cost and having the capability to alkylate benzene (or other aromatics) with the olefins would be beneficial to meet gasoline benzene specifications, increase motor fuel volume with high-octane alkylaromatic compounds and be economically acceptable in the current plant investment climate. For some refineries, the reactive removal of C2/C3 olefins could alleviate fuel gas capacity limitations. Such a process should:                Upgrade C2 and C3 olefin from fuel gas to high octane blending gasoline Increase flexibility in refinery operation to control benzene content in the gasoline blending pool        Allow refineries with benzene problems to feed the C6 components (low blending octane values) to the reformer, increasing both the hydrogen production from the reformer and the blend pool octane. Benzene produced in the reformer will be removed in order to comply with gasoline product specifications.        Have the potential, by the removal of olefins from the fuel gas, to increase capacity in the fuel system facility. For some refineries this benefit could allow an increase in severity in some key refinery process, FCC, hydrocracker, coker, etc.        
The necessity of keeping capital cost low obviously favors fixed bed catalytic units over the fluid bed type operations such as MOG and MBR. Fixed bed aromatics alkylation processes have achieved commercial scale use in the petrochemical field. The Cumene Process offered for license first by Mobil Oil Corporation and now by ExxonMobil Chemical Company is a low-capital cost process using a fixed bed of a zeolite alkylation/transalkylation catalyst to react refinery propylene with benzene to produce petrochemical grade cumene. Processes for cumene manufacture using various molecular sieve catalysts have been described in the patent literature: for example, U.S. Pat. No. 3,755,483 describes a process for making petrochemical cumene from refinery benzene and propylene using a fixed bed of ZSM-12 catalyst; U.S. Pat. No. 4,393,262 and U.S. also describe processes for making cumene from refinery benzene and propylene using ZSM-12 catalysts. The use of other molecular sieve catalysts for cumene manufacture has been described in other patents: U.S. Pat. No. 4,891,458 describes use of a zeolite beta catalyst; U.S. Pat. No. 5,149,894 describes the use of a catalyst containing the sieve material SSZ-25; U.S. Pat. No. 5,371,310 describes the use of a catalyst containing the sieve material MCM-49 in the transalkylation of diisopropyl benzene with benzene; U.S. Pat. No. 5,258,565 describes the use of a catalyst containing the sieve material MCM-36 to produce petrochemical grade cumene containing less than 500 ppm xylenes.
The petrochemical alkylation processes such as those referred to above, do not lend themselves directly to use in petroleum refineries without petrochemical capacity since they require pure feeds and their products are far more pure than required in fuels production. In addition, other problems may be encountered in the context of devising a process for motor gasoline production which commends itself for use in non-integrated, small-to-medium sized refineries. One such problem is the olefins from the cracker contain ethylene and propylene in addition to the higher olefins and if any process is to be economically attractive, it is necessary for it to consume both of the lightest olefins. Propylene is more reactive than ethylene and will form cumene by reaction with benzene at lower temperatures than ethylene will react to form ethylbenzene or xylenes (by transalkylation or disporportionation). Because of this, it is not possible with existing process technologies, to obtain comparable utilization of ethylene and propylene in a process using a mixed olefin feed from the FCCU. While improved ethylene utilization could in principle, be achieved by higher temperature operation, the thermodynamic equilibrium for the propylene/benzene reaction shifts away from cumene at temperatures above about 260° C. (500° F.), with consequent loss of this product.